Integrated Processes for Making Detergent Range Alkylbenzenes from C5-C6-Containing Feeds

ABSTRACT

Integrated processes for making detergent range alkylbenzenes from C 5 -C 6 -containing feeds involve feed pretreatment and/or selective hydrogenation to enable acceptable quality alkylbenzene production at attractive capital and operating costs.

CROSS-REFERENCE TO RELATED APPLICATION

This application is a Division of U.S. application Ser. No. 11/608,137filed Dec. 7, 2006, which claims benefit of U.S. Provisional ApplicationNo. 60/751,869, filed Dec. 20, 2005.

FIELD OF THE INVENTION

This processes disclosed herein pertain to integrated processes formaking alkylbenzenes suitable for use in detergent applications fromC₅-C₆-containing feeds such as derived from Fischer-Tropsch syntheses ornaphtha range petroleum fractions.

BACKGROUND TO THE INVENTION

Alkylbenzenes (phenyl-alkanes), especially where the alkyl group has 9to 14 carbon atoms, have found many utilities, the most prominent ofwhich is to make alkylbenzene sulfonates for use in laundry detergentsand similar products. Alkylbenzenes are prepared by the alkylation ofbenzenes with mono-olefins of the desired molecular weight andconfiguration of the sought alkyl group. Typically these olefins arederived from a kerosene range fraction from a petroleum refiningoperation. This fraction is a paraffin-containing fraction and issubjected to dehydrogenation to generate the sought olefins.

In some instances it may be desired to locate an alkylbenzene facilityat a location where a kerosene fraction is not readily available, orother commercial uses for a kerosene fraction render it lesseconomically attractive as a feedstock for making alkylbenzenes. Inthese instances the ability to use alternative feedstocks is highlydesirable.

Alternative feedstocks include other petroleum fractions, especiallynaphtha range fractions, and synthesized hydrocarbons such asFischer-Tropsch materials. These raw materials have a lower molecularweight than the sought olefins for alkylation and accordingly must besubjected to a dimerization or a metathesis to generate olefins ofsuitable chain length (detergent range olefins). Numerous processes havebeen disclosed for preparing detergent olefins from these alternativefeedstocks. See, for instance, WO 2004/072005, WO 2004/072006, and U.S.Patent Application Publications 2004/0030209, 2004/0176655, and2004/0199035.

In order for alternative feedstocks to be viable as a source ofdetergent range olefins, not only must the synthesized olefin meetconfigurational requirements such as degree and type of branching, butalso, the process must be technically feasible and economicallycompetitive. By way of illustration, US 2004/0176655 discloses that thepresence of oxygenated compounds and dienes is deleterious to theprocess and accordingly the feedstock is hydrogenated prior to adehydrogenation to generate olefins for a dimerization. While this stepeffectively removes oxygenates and dienes, it also hydrogenates anyolefins present in the feedstock and thus represents an operationalinefficiency. US 2004/0030209 discloses, inter alia, preparing an olefinmixture of pentene and hexene for dimerization by the metathesis of a C₄olefin mixture. Butadiene and acetylenic compounds are removed by aselective hydrogenation.

SUMMARY OF THE INVENTION

Process embodiments are provided for making alkylbenzenes from feedscontaining C₅ and C₆ paraffins that have enhanced operational economicsand yet provide alkylbenzenes of acceptable quality for commercial useincluding for sulfonation to produce detergents. Thus, alkylbenzenes canbe prepared from alternative, and often less expensive, feedstocks suchas naphtha range fractions and Fischer-Tropsch synthesized hydrocarbons.

In order to achieve these benefits, the processes integrate thepretreatment of the C₅ and C₆-containing feeds with the unit operationsfor generating detergent range olefins containing C₁₀ to C₁₂ olefins,and the production and refining of alkylbenzenes made from the detergentrange olefins. The integrated process embodiments use at least one ofthree integrated steps. One is removal of oxygenates from the C₅ andC₆-containing feeds using a lower alcohol or diol extractant to avoidthe necessity of a severe hydrogenation of the feed that not onlydestroys the oxygenates but also any existing olefinic unsaturation inthe feed. Thus, the second integration step is enabled which is tointroduce a feed containing C₅ and C₆ paraffins and olefins into a chaingrowth reaction zone (dimerization or metathesis which may be incombination with oligomerization) for producing the detergent rangeolefins prior to dehydrogenation of the paraffins in the feed. In thethird integration step, a selective hydrogenation is used to convertdiolefins to mono-olefins. By having removed the oxygenates, a selectivehydrogenation is sufficient to provide acceptable alkylbenzene product.The selective hydrogenation may occur between a dehydrogenation toproduce olefins from the feed containing C₅ and C₆ paraffins and theformation of the detergent range olefins, or the selective hydrogenationmay occur after the formation of the detergent range olefins. Aparticularly attractive aspect of the process embodiments is that by theselective hydrogenation of the dehydrogenation product, thedehydrogenation can be conducted at higher conversion conditions thatinherently also increase the formation of diolefins. Moreover, thispreferred aspect of the process embodiments facilitates the refining ofthe alkylbenzene product since virtually no paraffins in the C₁₀ to C₁₂range are generated. These higher molecular weight paraffins wouldtypically require a dedicated distillation column operating atsubatmospheric pressure for removal from alkylbenzenes.

One of the broad aspects of the integrated process embodiments forproducing detergent range alkylbenzenes from paraffinic feedstockcomprises:

a) subjecting a feedstock predominantly comprising C₅ and C₆ paraffinsto dehydrogenation conditions to provide a dehydrogenated product thatcomprises C₅ and C₆ mono-olefins and a minor amount of diene;

b) reacting the dehydrogenated product of step a under chain growthconditions to provide a detergent range olefin product comprising C₁₀ toC₁₂ mono-olefins;

c) selectively hydrogenating at least one of the dehydrogenated productof step a and the detergent range olefin product of step b to reducediene content;

d) alkylating benzene with at least a portion of the detergent rangeolefin product having been subjected to selective hydrogenation of stepc under alkylation conditions including a stoichiometric excess ofbenzene to olefin to provide an alkylation effluent comprisingalkylbenzene and benzene; and

e) separating benzene from the alkylation effluent from the alkylbenzeneand recycling at least a portion of the benzene to step d.

Another broad aspect of the integrated process embodiments for producingdetergent range alkylbenzenes from paraffinic feedstock comprises:

a) subjecting a feedstock predominantly comprising C₅ and C₆ paraffinsand containing C₅ and C₆ mono-olefins to chain growth reactionconditions for converting C₅ and C₆ mono-olefins to C₁₀ to C₁₂mono-olefins to provide a reaction product containing C₅ and C₆paraffins and C₁₀ to C₁₂ mono-olefins;

b) subjecting the reaction product to distillation to provide a higherboiling fraction containing C₁₀ to C₁₂ mono-olefins and a lower boilingfraction comprising C₅ and C₆ paraffins;

c) subjecting at least a portion of the lower boiling fraction of step bto dehydrogenation conditions to provide a dehydrogenated productcomprising C₅ and C₆ mono-olefins;

d) recycling at least a portion of the dehydrogenated product to step a;

e) alkylating benzene with at least a portion of the higher boilingfraction of step b under alkylation conditions including astoichiometric excess of benzene to olefin to provide an alkylationeffluent comprising alkylbenzene and benzene; and

f) separating benzene from the alkylation effluent from the alkylbenzeneand recycling at least a portion of the benzene to step e.

Yet another broad aspect of the integrated process embodiments forproducing detergent range alkylbenzenes from paraffinic feedstockcomprises:

a) contacting a feedstock predominantly comprising C₅ and C₆ paraffinsand oxygenated compounds with a liquid extractant comprising at leastone of alcohol or diol of 1 to 3 carbon atoms per molecule and a minoramount of water under extraction conditions to provide a deoxygenatedfeedstock having a reduced concentration of oxygenated compounds andspent extractant;

b) regenerating and recycling to step a at least a portion of the spentextractant as liquid extractant;

c) dehydrogenating the deoxygenated feedstock in a dehydrogenation zoneand recovering therefrom a dehydrogenated product comprising C₅ and C₆mono-olefins;

d) reacting the dehydrogenated product of step a under chain growthconditions to provide a detergent range olefin product comprising C₁₀ toC₁₂ mono-olefins;

e) alkylating benzene with at least a portion of the detergent rangeolefin product under alkylation conditions including a stoichiometricexcess of benzene to olefin to provide an alkylation effluent comprisingalkylbenzene and benzene; and

f) separating benzene from the alkylation effluent from the alkylbenzeneand recycling at least a portion of the benzene to step d.

In preferred aspects of the process embodiments, the detergent rangeolefin product used for benzene alkylation is substantially free ofparaffins having 9 or more carbon atoms per molecule, e.g., less thanabout 0.5, preferably less than about 0.1, mass-percent of paraffinshaving 9 or more carbon atoms per molecule is present in the detergentrange olefin. Thus, the separation of alkylbenzene from the alkylationeffluent can be effected in a cost effective manner. Where paraffinshaving 9 or more carbon atoms per molecule are present, processes forrecovering alkylbenzenes have relied upon a paraffin distillation columnfor purification of the alkylbenzenes. Moreover, these distillationcolumns are typically operated at subatmospheric pressures, often belowabout 20 kPa(absolute) (hereinafter kPa(a)) (2.9 psi(absolute)(hereinafter psi(a))), to avoid temperatures that can result indegradation of the alkylbenzene product. These preferred aspects of thisembodiment, by not requiring a separation of paraffins having 9 or morecarbon atoms per molecule from the alkylbenzene product, provide forreduced capital and operating costs.

In another preferred aspect of the process embodiments, the alkylationis conducted in the presence of a paraffinic component having between 4and 7, preferably 5 to 6, carbon atoms per molecule to assist inmoderating the exothermic alkylation reaction. Preferably, theparaffinic component is provided to the alkylation reaction in a moleratio to detergent range olefin of between about 1:1 and 20:1. Theparaffinic component is readily separated from the alkylbenzene bydistillation to provide a benzene and paraffin-containing fraction thatcan be recycled to the alkylation reaction.

In other preferred aspects of this embodiment, the dehydrogenation ofthe feed containing C₅ and C₆ paraffins, is conducted at higherconversions to olefins. Higher conversions generate greater amounts ofdienes and cyclics. Preferred integrated processes can tolerate thegreater concentrations of dienes and cyclics in the dehydrogenationeffluent due to the selective hydrogenation and the separation of thedetergent range olefins from lower boiling components which includecyclics. Cyclics generally comprise benzene with some cyclopentadiene,naphthenes and toluene possibly also being formed. The separateddetergent range olefins preferably contain less than about 0.1, morepreferably less than about 0.01, mass-percent cyclics. Advantageously,benzene can be recovered to provide at least a portion of the benzenefor the alkylation.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic depiction of an apparatus for making alkylbenzenesin which a synthesized C₅-C₆-containing feed is dimerized and thedimerized feed is used to supplement a detergent-range olefin stream asa feed to an alkylation reactor.

FIG. 2 is a schematic depiction of an apparatus for making alkylbenzenesin which a C₅-C₆ feed is metathesized to provide a detergent-rangeolefin stream as a feed to an alkylation reactor.

FIG. 3 is a schematic depiction of an apparatus for making alkylbenzenesfrom a C₅-C₆ feed with the co-production of benzene duringdehydrogenation for use in the alkylation.

DETAILED DESCRIPTION

The process embodiments relate to the manufacture of alkylbenzenessuitable for detergent applications from lower molecular weightparaffins, i.e., feedstocks predominantly comprising C₅ and C₆paraffins. One source for such feedstocks is from naphtha rangefractions from petroleum refining. As branched paraffins and naphthenesare generally preferred for gasoline, normal C₅ and C₆ paraffinstypically have less value to the refiner. However, they tend to bepreferred feeds for alkylbenzene manufacture due to the linearity soughtfor alkylbenzenes used in making detergents. Thus, while the processembodiments can use naphtha range fractions, advantageously a portion ofthe naphtha fraction is used which contains a greater concentration ofnormals. This preferred feed may thus be derived from a separation ofthe naphtha range fraction or may be from a process unit operationinvolved in gasoline refining such as the effluent from an isomerizationunit. Thus, a benefit of the process embodiments can be the up-gradingof naphtha range fractions by selectively removing the low octane normalhydrocarbons for use in making detergent range olefins. The separationof normal hydrocarbons from naphtha range fractions is described in, forinstance, Stephen W. Sohn, UOP Molex Process for Production of NormalParaffins, Handbook of Petroleum Refining Processes, Second Edition,Editor Robert A. Meyers, McGraw-Hill, New York, USA (1997) at pp. 10.75to 10.77. Another source of feedstocks predominantly comprising C₅ andC₆ paraffins is from synthesis processes such as a Fischer-Tropschprocess.

While the feedstocks used in the process embodiments will have differentcompositions depending upon the source of the feedstocks, the feedstockwill nevertheless be predominantly composed of C₅ and C₆ paraffins. Asused herein, a feedstock is predominantly composed of a material whenthe feedstock contains about 50 mass-percent or more of the material.The feedstocks will likely also contain hydrocarbons of higher and lowercarbon numbers. Typically the feedstocks will have the followingcompositions:

Component Typical, mass-percent Preferred, mass-percent C₄ 0.0 to 10 0.0to 5 C₅ and C₆ paraffins 50 to 100 70 to 100 C₅ and C₆ olefins 0 to 50 0to 30 C₇+ 0 to 20 0 to 10 Aromatics 0 to 10 0 to 5 Oxygenates 0 to 3 0to 1 Dienes 0 to 2 0 to 1

In the broad aspects of this embodiment wherein the feedstock containsboth C₅ and C₆ paraffins and C₅ and C₆ mono-olefins and is fed to adimerization or metathesis reactor prior to dehydrogenation, the olefinsgenerally comprise at least about 1, and preferably about 1 to 50, oftenbetween about 15 and 30, and in some cases from about 2 to 10,mass-percent of the feedstock. The feedstock may be combined with atleast a portion of the product of a dehydrogenation prior to beingsubjected to chain growth conditions for dimerization or metathesis, orit may be subjected to chain growth conditions to convert at least aportion of the olefinic component to higher carbon number olefins priorto combination with product from the dehydrogenation. In either event,at least a portion of the olefins in the feedstock are consumed and notpassed to the dehydrogenation. Not only is a portion of the olefins inthe feedstock effectively used, but also, the concentration of olefinsin the stream to be dehydrogenated is reduced. As the dehydrogenation iseffected by equilibria, enhanced efficiencies in net conversion ofparaffins to olefins can be realized.

In the broad aspects of this embodiment where the feedstock containsoxygenated components such as alcohols, aldehydes, ketones, ethers,acids and esters, the oxygenated components may be present in amounts upto about 10 mass-percent. Normally, the oxygenates are present in therange of about 0.001 to 10 mass-percent. More than one oxygenate may bepresent. The extraction in accordance with this aspect of the embodimentoften will yield a feedstock containing less than about 1000, andpreferably less than about 200, and sometimes even below 20, parts permillion by mass (ppm-mass) oxygenates. See, for instance, Jansen, etal., “Novel Approach Using Methanol/Water Extraction forHydrocarbon/Oxygenate Separation”, presented at the ExtractiveSeparations Topical Conference AIChE Annual Meeting, San Francisco,Calif., November, 2003, for background on extractive separationprocesses for removing oxygenates from paraffins and olefins.

The extraction is effected using a liquid extractant comprising at leastone of alcohol or diol of 1 to 3 carbon atoms per molecule and a minoramount of water under extraction conditions. As used herein, a fluid hasa minor amount of a material when the fluid contains about 25mass-percent or less of the material. The alcohols and diols may be oneor more of ethanol, 1-propanol, 2-propanol, ethylene glycol, propyleneglycol and 1,3-propanediol, and preferably methanol. The amount of waterpresent in the extractant is usually less than about 25 mass-percentalthough more water can be present, little benefit is normally seen inusing the higher concentrations of water. Most often water is present inthe range of about 2 to 15, say, 3 to 10, mass-percent.

Suitable extraction conditions maintain the extractant and the feedstockin liquid phase during the extraction. Usually the extraction isconducted at temperatures in the range of from about 15° (59° F.) to150° C. (302° F.), preferably 30° (86° F.) to 120° C. (258° F.), andpressures of from about 100 kPa(a) (14.5 psi(a)) to 5000 kpa(a) (725psi(a)), frequently, from about 150 kPa(a) (21.8 psi(a)) to 1000 kpa(a)(145 psi(a)). The extraction operation may be operated in any convenientmanner. Normally the contact time is between about 5 seconds and 10minutes as sufficient to provide the desired reduction in oxygenatecontent. The extraction may be conducted in a liquid-liquid extractioncolumn. The column may, if desired, contain packing to assist inliquid-liquid contact. The extraction can also be effected in a vesselin which the liquid streams are agitated.

The feedstock is contained in the raffinate and can be recovered by anysuitable means such as extraction with water or sorption, can be used toremove residual alcohol or diol from the feedstock. Normally thefeedstock will contain less than about 20,000 mass-ppm alcohol or diolafter such stripping or sorption. The solvent phase from the extractionwill contain in addition to the alcohol or diol water and oxygenates.The alcohol or diol can be separated for recycle, e.g., by stripping orfractionation of the solvent phase. The oxygenates pass with the waterduring such separations.

Typically oxygenates are removed from the feedstock prior todehydrogenation. Since many dehydrogenation processes can tolerate thepresence of water and many types of oxygenates such as alcohols,aldehydes, ketones and carboxylic acids that convert to water underdehydrogenation conditions, the dehydrogenation may alternatively beconducted prior to the removal of oxygenates.

Dehydrogenation is conducted in the presence of hydrogen anddehydrogenation catalyst under dehydrogenation conditions includingelevated temperature. Any suitable dehydrogenation system may be used.In general, dehydrogenation conditions are selected to minimize crackingand polyolefin by-products. It is expected that typical dehydrogenationconditions will not result in any appreciable skeletal isomerization ofthe hydrocarbons in the dehydrogenation reactor. Dehydrogenationconditions include a temperature of generally from about 400° C. (752°F.) to about 900° C. (1652° F.) and preferably from about 420° C. (788°F.) to about 550° C. (1022° F.), a pressure of generally from about 1kpa(g) (0.15 psi(g)) to about 1000 kPa(g) (145 psi(g)), and a LHSV offrom about 0.1 to about 100 hr⁻¹. As used herein, the abbreviation“LHSV” means liquid hourly space velocity, which is defined as thevolumetric flow rate of liquid per hour divided by the catalyst volume,where the liquid volume and the catalyst volume are in the samevolumetric units. Generally for normal paraffins, the lower themolecular weight the higher the temperature required for comparableconversion. The pressure in the dehydrogenation zone is maintained aslow as practicable, usually less than 350 kPa(g) (50.8 psi(g)) tomaximize chemical equilibrium advantages.

The feedstock containing C₅ and C₆ paraffins may be admixed with adiluent material before, while, or after passing to the dehydrogenationzone. The diluent material may be hydrogen, steam, methane, ethane,carbon dioxide, nitrogen, argon, and the like, or a mixture thereof.Hydrogen is the preferred diluent. Ordinarily, when hydrogen is utilizedas the diluent it is utilized in amounts sufficient to ensure a hydrogento hydrocarbon mole ratio of about 0.1:1 to about 40:1, with bestresults being obtained when the mole ratio range is about 1:1 to about10:1. The diluent hydrogen stream passed to the dehydrogenation zonewill typically be recycled hydrogen separated from the effluent from thedehydrogenation zone in the hydrogen separation zone.

Water or a material which decomposes at dehydrogenation conditions toform water such as an alcohol, aldehyde, ether, or ketone, for example,may be added to the dehydrogenation zone, either continuously orintermittently, in an amount to provide, calculated on the basis ofequivalent water, about 1 to about 20,000, preferably about 1 to about10,000, mass-ppm of the hydrocarbon feed stream.

Any suitable dehydrogenation catalysts can be used. The choice of aparticular dehydrogenation catalyst is not critical to this embodiment.Typically the catalyst comprises at least one platinum group (Groups8-10 of the periodic table) metal and at least one promoter metal on asuitable support, usually a refractory oxide such as alpha alumina,theta alumina, cordierite, zirconia, titania, and mixtures thereof.Platinum group metals include platinum, palladium, rhodium, iridium,ruthenium, and osmium. Promoter metals typically are selected from thegroup consisting of tin, germanium, rhenium, gallium, bismuth, lead,indium, cerium, silver, zinc, and mixtures thereof, while modifiermetals are selected from the group consisting of alkali metals, alkalineearth metals such as lithium, sodium, potassium, cesium, rubidium,beryllium, magnesium, calcium, strontium, barium, and mixtures thereof.The concentration of each metal component can vary substantially. Theplatinum group metal is often present in a concentration of about 0.01to about 5 mass-percent on an elemental basis of the entire mass of thecatalyst. The promoter metal may be present in an amount from about 0.05to about 5 mass-percent while the modifier metal may be present in anamount from about 0.1 to about 5 mass-percent. The atomic ratio of theplatinum group metal to promoter metal typically falls within the rangeof from about 0.05 to about 5. See, for instance, U.S. Pat. Nos.3,274,287; 3,315,007; 3,315,008; 3,745,112; 4,430,517; 4,716,143;4,762,960; 4,786,625; and 4,827,072.

The dehydrogenation catalyst may be in a fixed bed or a moving catalystbed or a fluidized bed. The dehydrogenation zone may comprise one ormore catalyst-containing reaction zones with heat exchangers therebetween to ensure that the desired reaction temperature is maintained atthe entrance to each reaction zone. One or more hot hydrogen-rich gasstreams may be introduced between a first and a second reaction zone toincrease the temperature of a stream passing from the first to thesecond reaction zone.

Each reaction zone may be operated in a continuous-type or batch-typemanner. Each reaction zone may contain one or more catalyst beds.Hydrocarbons may contact any catalyst bed in an upward-, downward-, orradial-flow fashion. In a particularly compact and efficientarrangement, the contacting of the catalyst with hydrocarbons and heatexchanging may be accomplished in a heat exchanging reactor.

The dehydrogenated product is typically a mixture of unreacted paraffinsand olefins. The conversion of paraffins to olefins is equilibriumlimited. Normally, the conversion of paraffins to olefins is betweenabout 5 and 40, preferably between about 7 and 30, mole percent. Withlimited conversion of paraffins to olefins, it may be possible togenerate little dienes. See, for instance, US 2004/0176655 where a lowconversion per pass dehydrogenation unit operation is used to avoidgeneration of deleterious amounts of dienes.

In the preferred aspects of the process embodiments, however, thedehydrogenation conditions are sufficiently severe that higherconversions of paraffins to olefins are obtained, e.g., from about 10 to35, mole percent even though these conditions result in diene beingco-produced. Typically under these more severe conditions, the diene ispresent in an amount of at least about 1, say, about 3 to 25,mass-percent based upon the mass of the total olefin (mono- anddi-olefin) in the dehydrogenated product. Also, in this preferredembodiment, cyclics are formed. As the feedstock generally containslittle C₇ and higher hydrocarbons, the cyclics are rich in benzene.Normally, of the cyclics, benzene constitutes at least about 40, say,between about 40 and 95, mass-percent of the cyclics. The conversion ofparaffins to cyclics is often in an amount between about 1 and 10, say,about 5 and 20, mass-percent of the paraffins fed to thedehydrogenation. In one embodiment, the concentration of cyclics oraromatic compounds in the dehydrogenated product is at least 5mass-percent. The concentration of cyclics or aromatic compounds in thedehydrogenated product is often between about 1 and 10, say, about 5 and20 mass-percent. The concentration of cyclics in the dehydrogenatedproduct may be greater as unreacted hydrocarbons may be recycled to thedehydrogenation step and cyclics can build-up in the recycle stream.

The dehydrogenated product is the feed for synthesizing detergent rangeolefins. Especially where the dehydrogenation has been conducted toachieve higher conversions of paraffins, and thus contains diene, thedehydrogenated product is subjected to selective hydrogenation. If thedehydrogenation is under low conversion per pass conditions, it may beacceptable only to subject the detergent olefin-containing productstream to selective hydrogenation.

In broad terms, regardless of whether the dehydrogenated product or thedetergent olefin-containing product stream is subjected to selectivehydrogenation, the stream to be selectively hydrogenated is contactedwith selective hydrogenation catalyst under selective hydrogenationconditions including the presence of hydrogen and elevated temperature.Selective hydrogenation processes are normally performed at relativelymild hydrogenation conditions. A broad range of suitable operatingpressures therefore extends from about 40 kPa(g) (5.8 psi(g)) to about250 kPa(g) (36.3 psi(g)) to 6000 kpa(g) (870 psi(g)), preferably betweenabout 300 kPa(g) (43.5 psi(g)) and 2000 kPa (g) (290 psi(g)). Arelatively moderate temperature between about 25° C. (77° F.) to about350° C. (682° F.), preferably 50° C. (122° F.) to 200° C. (392° F.), isoften used. The mole ratio of hydrogen to diolefinic hydrocarbons toachieve a certain conversion is dependent upon both reactor temperatureand the molecular weight of the feed hydrocarbons. To avoid theundesired saturation of a significant amount mono-olefinic hydrocarbons,there should be less than 2.0 times the stoichiometric amount ofhydrogen required for the selective hydrogenation of the diolefinichydrocarbons which are present in the liquid phase process stream tomono-olefinic hydrocarbons. Preferably, the mole ratio of hydrogen todiolefinic hydrocarbons in the material entering the bed of selectivehydrogenation catalyst is maintained between 0.75:1 and 1.8:1. Theselective hydrogenation catalyst may be any suitable catalyst foreffecting mild hydrogenation, and are usually supported hydrogenationmetal such as nickel (especially sulfided nickel), platinum orpalladium. See, for instance, U.S. Pat. Nos. 3,234,298; 3,472,763;3,662,015; 4,695,560 and 5,276,231, all of which are herein incorporatedin their entireties by reference. The liquid hourly space velocity ofthe reactants through the selective hydrogenation catalyst should beabove 1.0 hr⁻¹. Preferably, it is above 5 hr⁻¹, and more preferably itis between 5 and 35 hr⁻¹. In preferred aspects, the selectivehydrogenation conditions are sufficient to reduce the dieneconcentration to less than about 0.5, preferably less than about 0.3,mass-percent dienes based upon the total mass of olefins in theselective hydrogenation effluent.

Where the selective hydrogenation follows the chain growth reaction toform the detergent range olefins, the dehydrogenation is preferablyconducted to minimize the formation of dienes. Typically, the detergentrange olefins would be separated from the lower boiling hydrocarbons andthis separated stream would be subjected to the dehydrogenationconditions. This detergent range olefin-containing would contain similarcarbon atom-containing dienes, but, because of the less severedehydrogenation conditions employed, in a relatively small amount, e.g.,less than about 1, preferably less than about 0.5, mass-percent basedupon the total olefins.

Alternatively, especially where the feedstock containing C₅ and C₆paraffins also contains olefins and, prior to dehydrogenation, issubjected to chain growth reaction conditions to form a detergent rangeolefin product comprising C₁₀ to C₁₂ mono-olefins, the feedstock may beselectively hydrogenated. In this embodiment, the feedstock ispreferably admixed with the recycling lower boiling fraction ofparaffins and unreacted olefins from the reaction to form the detergentrange olefin product. This admixture is subjected to the selectivehydrogenation to reduce the presence of any dienes that may be in thefeedstock.

The selective hydrogenation catalyst may be in a fixed bed or a movingcatalyst bed or a fluidized bed. The selective hydrogenation zone maycomprise one or more catalyst-containing reaction zones with heatexchangers there between to ensure that the desired reaction temperatureis maintained at the entrance to each reaction zone. One or morehydrogen-rich gas streams may be introduced between a first and a secondreaction zone to increase the temperature of a stream passing from thefirst to the second reaction zone.

Each reaction zone may be operated in a continuous-type or batch-typemanner. Each reaction zone may contain one or more catalyst beds.Hydrocarbons may contact any catalyst bed in an upward-, downward-, orradial-flow fashion. In a particularly compact and efficientarrangement, the contacting of the catalyst with hydrocarbons and heatexchanging may be accomplished in a heat exchanging reactor.

The dehydrogenated product which contains C₅ and C₆ olefins is subjectedto chain growth reaction conditions sufficient to provide a detergentrange olefin product comprising C₁₀ to C₁₂ mono-olefins (herein referredto as the “chain growth reaction”). The particular type of chain growthreaction is not critical to the broadest aspects of the processembodiments. The chain growth reaction may be one or more ofdimerization and metathesis, which may be in combination witholigomerization.

Dimerization conditions typically include the presence of a catalyst andthe use of elevated temperature and pressures. The specific temperatureand pressure conditions used will depend, in part, upon the type ofcatalyst employed. A wide variety of catalysts have been proposedincluding homogeneous and heterogeneous catalysts. Examples ofhomogeneous catalysts include hydrogen fluoride, boron trifluoride andtrifluoroacetic acid. Heterogeneous catalysts include suitablesilica-aluminas, sulfated zirconias and molecular sieves and supportedmetal-containing catalysts that often contain at least one elementselected from Groups 3, 4, 8 to 10, and 14 of the periodic table.References herein to the periodic table are to the new IUPAC notation asshown on the Periodic Table of the Elements in the inside front cover ofthe CRC Handbook of Chemistry and Physics, 80th Edition, 1999-2000, CRCPress, Boca Raton, Fla., USA. The metal element is usually in combinedform such as an oxide, sulfide or oxysulfide. Due to selectivity todimerization, catalysts containing nickel oxide or iron oxide are mostfrequently used. Suitable supports include silica, alumina,silica-aluminas including molecular sieves. See, for instance, U.S. Pat.Nos. 5,169,824 and 5,849,972. Temperatures for the dimerization aregenerally in the range of about 40° C. (104° F.) to 250° C. (482° F.),preferably about 60° C. (140° F.) to 200° C. (392° F.), and pressuresmay be within range of 100 kpa(a) (14.5 psi(a)) to 2000 kpa(a) (290psi(a)), say, 110 kpa(a) (16 psi(a)) to 1000 kpa(a) (145 psi(a)).

Heterogeneous dimerization catalyst may be in a fixed bed or a movingcatalyst bed or a fluidized bed. The dimerization zone may comprise oneor more catalyst-containing reaction zones with heat exchangers therebetween to ensure that the desired reaction temperature is maintained atthe entrance to each reaction zone. Preferably more than one reactionzone is used with intervening fractionation to remove dimerized product.Each reaction zone may be operated in a continuous-type or batch-typemanner. Each reaction zone may contain one or more catalyst beds.Hydrocarbons may contact any catalyst bed in an upward-, downward-, orradial-flow fashion.

Preferably the dimerization conditions are such that on a per pass basisat least about 20, more preferably about 30 to 70, mass-percent of theolefin is converted to dimers and highers with at least about 70,preferably at least about 75, mass-percent of the conversion being todimers. Preferably little, if any, C₁₀ to C₁₂ paraffins are formedduring the dimerization, and often the dimerization product containsless than about 0.1, preferably less than about 0.01, mass-percent C₁₀to C₁₂ paraffins.

Metathesis reaction conditions include the presence of a metathesiscatalyst. The catalyst may be homogeneous or heterogeneous.Catalytically-active elements proposed for metathesis catalysts includeelements from Groups 4, 5, 6, 7, and 8 to 10 of the periodic table,including one or more of titanium, niobium, tantalum, molybdenum,tungsten, rhenium, ruthenium, osmium and iridium, especially one or moreof rhenium, tungsten and molybdenum. The ultimate form of thecatalytically active component may be an oxide, sulfide or oxysulfide orin a ligand structure. The heterogeneous catalysts are generallysupported, for instance, on a refractory oxide support or a molecularsieve-containing support. See, for instance, Chapter 2 of “OlefinMetathesis and Metathesis Polymerization” by K. J. Ivin and J. C. Mol,Academic Press, San Diego, Calif., USA (1997). Temperatures for themetathesis are generally in the range of about 20° C. (68° F.) to 300°C. (572° F.), preferably about 35° C. (95° F.) to 150° C. (302° F.), andpressures may be within range of 100 kPa(a) (14.5 psi(a)) to 2000 kpa(a)(290 psi(a)), say, 110 kpa(a) (16 psi(a)) to 1000 kPa(a) (145 psi(a)).

The heterogeneous metathesis catalyst may be in a fixed bed or a movingcatalyst bed or a fluidized bed. The metathesis zone may comprise one ormore catalyst-containing reaction zones. Each reaction zone may beoperated in a continuous-type or batch-type manner. Each reaction zonemay contain one or more catalyst beds. Hydrocarbons may contact anycatalyst bed in an upward-, downward-, or radial-flow fashion. Each zonemay have associated with it a fractionation column to recover andrecycle lower molecular weight olefin and remove detergent range olefinproduct.

Preferably the metathesis conditions are such that at least about 30,more preferably about 40 to 95, mass-percent of the olefin is convertedto detergent range olefins. Preferably little, if any, C₁₀ to C₁₂paraffins are formed during the metathesis, and often the metathesisproduct contains less than about 0.1, preferably less than about 0.01,mass-percent C₁₀ to C₁₂ paraffins.

The by-products of metathesis are ethylene and propylene that havevalue, especially in an integrated refinery and alkylbenzene productionfacility.

Oligomerization can be used where the C₅-C₆ feedstock also contains asignificant amount of C₄ hydrocarbons or in combination with ametathesis chain growth operation to convert lower olefins to detergentrange olefins or to C₅ and C₆ olefins for metathesis or dimerization.Oligomerization conditions typically include the presence of a catalystand the use of elevated temperature and pressures. The specifictemperature and pressure conditions used will depend, in part, upon thetype of catalyst employed. A wide variety of catalysts have beenproposed including homogeneous and heterogeneous catalysts. Examples ofhomogeneous catalysts include hydrogen fluoride, boron trifluoride andtrifluoroacetic acid. Homogeneous catalysts also include metalcoordination catalysts such as Ziegler-Natta and metallocene-typecatalysts. Heterogeneous catalysts include suitable silica-aluminas,sulfated zirconias and molecular sieves and supported metal-containingcatalysts. The metals for the homogeneous and heterogeneous catalystsoften comprise at least one element selected from Groups 3, 4, 14, and 8to 10 of the periodic table. The element is usually in combined formsuch as an oxide, sulfide or oxysulfide or in a ligand structure.Catalysts containing nickel, cobalt or iron, usually as oxides, are mostfrequently used. Suitable supports include silica, alumina,silica-aluminas including molecular sieves. See, for instance, U.S. Pat.Nos. 5,169,824 and 5,849,972. Temperatures for the oligomerization aregenerally in the range of about 100° C. (212° F.) to 250° C. (482° F.),preferably about 120° C. (248° F.) to 200° C. (392° F.), and pressuresmay be within range of 100 kPa(a) (14.5 psi(a)) to 2000 kPa(a) (290psi(a)), say, 110 kPa(a) (16 psi(a)) to 1000 kPa(a) (145 psi(a)).

The heterogeneous oligomerization catalyst may be in a fixed bed or amoving catalyst bed or a fluidized bed. The oligomerization zone maycomprise one or more catalyst-containing reaction zones with heatexchangers there between to ensure that the desired reaction temperatureis maintained at the entrance to each reaction zone. Preferably morethan one reaction zone is used with intervening fractionation to removeoligomerized product. Each reaction zone may be operated in acontinuous-type or batch-type manner. Each reaction zone may contain oneor more catalyst beds. Hydrocarbons may contact any catalyst bed in anupward-, downward-, or radial-flow fashion.

Preferably the oligomerization conditions are such that on a per passbasis at least about 20, more preferably about 30 to 70, mass-percent ofthe olefin is converted to dimers and highers with at least about 60,preferably at least about 70, mass-percent of the conversion being todetergent range olefins. Preferably little, if any, C₁₀ to C₁₂ paraffinsare formed during the oligomerization, and often the product containsless than about 0.1, preferably less than about 0.01, mass-percent C₁₀to C₁₂ paraffins.

In preferred embodiments, a detergent range olefin fraction isseparated, usually by distillation, from the effluent from the chaingrowth reaction. A lower molecular weight fraction, e.g., having C₄-C₇hydrocarbons, can be recycled to the dehydrogenation step. As thedehydrogenated product contains significant amounts of C₅ and C₆paraffins due to the equilibrium limitation on conversion, a recycleprovides for efficient utilization of feedstock. Normally, this recyclestream will contain from about 40 to 98 mass-percent paraffins, lessthan 5 mass-percent olefins, and, depending upon the co-production ofcyclics during the dehydrogenation and the volume of the cyclics removedfrom the reaction loop discussed below, from about 0.1 to 60,mass-percent cyclics.

To prevent an undue build-up of cyclics (i.e., above 60 mass-percent) ina reaction loop with the dehydrogenation and the chain growth reaction,a cyclics-containing stream is typically removed. The removal may beeffected by subjecting all or a portion of the fluid in the reactionloop to a separation process to remove cyclics. While distillation canbe used, membrane separation, extraction and sorption techniques aregenerally preferred for selective removal of cyclics, especiallyaromatics. In a preferred aspect of the process embodiments, a slipstream, or purge, is taken from the loop, preferably downstream from thechain growth reaction.

Benzene is a primary constituent of the cyclics and is a raw materialfor making alkylbenzenes. Where the dehydrogenation is operated to formaromatics, preferably at least a portion of the benzene is recovered andused as a portion of the benzene feed for the alkylation step. Becausethe process embodiments use a feed primarily containing C₅ and C₆paraffins, the formation of higher, less desirable aromatics such astoluene, is much less than that occurring with the dehydrogenation of aC₁₀ to C₁₂ paraffin-containing stream.

In one embodiment, at least a portion of the purge is subjected todealkylation conditions to convert any toluene and higher aromatics tobenzene. If the dealkylation is done under hydrocracking conditions, thedealkylation conditions may also hydrogenate any olefins present. Thebenzene can be separated from the dealkylation effluent or thedealkylation effluent can be fed to the alkylation reaction, and thenon-aromatics in admixture with the benzene can serve as a heat sink toassist in removal of heat from the exothermic alkylation reaction.Alternatively, at least a portion of the non-aromatics can be separatedand, if desired, recycled to the dehydrogenation.

Suitable dealkylation processes for benzene-forming includehydrodealkylation, steam dealkylation, oxidative dealkylation incombination with steam or among others. Hydrodealkylation can be thermaland or catalytic. Thermal hydrodealkylation is preferred. Thebenzene-forming conditions are preferably sufficient to dealkylate atleast about 70, more preferably at least about 85, and sometimes betweenabout 85 and 98, mass-percent of the aromatic by-products.

In general, hydrodealkylation conditions include elevated temperaturesand pressures, e.g., about 200° C. (392° F.) to 700° C. (1292° F.) and100 kPa(g) (14.5 psi(g)) to 5000 kPa(g) (725 psi(g)) and the presence ofhydrogen, e.g., from about 1:1 to 50:1 moles of hydrogen per mole ofaromatic by-products. Thermal hydrodealkylation is typically attemperatures of between about 500° C. (932° F.) or 575° C. (1067° F.)and 700° C. (1291° F.). Often the pressures of thermal hydrodealkylationare in the range of about 2500 kPa(g) (363 psi(g)) to 3500 kpa(g) (508psi(g)). Exemplary thermal hydrodealkylation conditions are describedin, for instance, W. L. Liggin, UOP Thermal Hydrodealkylation (THDA)Process, in Handbook of Petroleum Refining Processes, Second Edition(1997), at pp. 2.23 to 2.26.

Catalytic hydrodealkylation conditions are generally milder than thoseused for thermal hydrodealkylation. Temperatures of about 200° C. (392°F.) to 600° C. (1112° F.), say, 250° C. (482° F.) to 550° C. (1022° F.),are often employed. Pressures are generally in the range of about 100kPa(g) (14.5 psi(g)) to 3500 kpa(g) (508 psi(g)). Often, the mildercatalytic hydrodealkylation conditions will result in less cracking ofthe alkyl groups removed from the aromatic by-products, that is, lightends may contain a lesser concentration of methane than would occur froma thermal hydrodealkylation. Hence, less hydrogen may be required, e.g.,from about 1 to 10 moles of hydrogen per mole of aromatic by-products.Any suitable hydrodealkylation catalyst may be used. Typical catalystsare acidic catalysts including aluminas, silicas, silica aluminas suchas zeolites including dealuminated Y, silicalite, zeolites having theMOR framework type, zeolite beta, ZSM-5, zeolites having the MTWframework type, UZM-4, UZM-5, UZM-8, UZM-16; silica aluminophosphatessuch as SAPO-34, SAPO-11, SAPO-31, SAPO-5, SAPO-18, and MAPSO-43; whichmay be supported or unsupported. References herein to zeolite frameworktypes are to those in the Atlas of Zeolite Framework Types, FifthRevised Edition, 2001, Elsevier, N.Y., USA. UZM-4 is described in U.S.Pat. No. 6,419,895, herein incorporated by reference in its entirety.UZM-5 and UZM-6 are described in U.S. Pat. Nos. 6,613,302 and 6,388,157,which are herein incorporated in their entireties by reference. UZM-8 isdescribed in U.S. Pat. No. 6,756,030, herein incorporated by referencein its entirety. UZM-16 is described in U.S. Pat. No. 6,752,980, hereinincorporated by reference in its entirety. In some instances, thecatalyst is halogenated, e.g., fluorided or chlorided, to enhanceacidity. The catalyst may also comprise one or more adjuvants or agentswhich may serve as catalysts, promoters and activity modifiers. Activemetals can be from Groups 4, 5, 6, 7, and 8 to 10 of the periodic table,such as platinum and rhenium. Promoters and/or activity modifiers can beselected from Groups 1, 2, 13, 14, 15, and 16 of the periodic table,such as tin, lead, germanium, and sulfur.

As the benzene-forming conditions typically involve the presence ofhydrogen, preferred process embodiments recover unreacted hydrogen fromthe product of the benzene-forming conditions for recycle. Generallythis recovery is done by gas/liquid phase separation.

A benzene product stream is generated as a result of subjecting thecyclics-containing stream to benzene-forming conditions. This benzeneproduct stream will contain benzene and non-aromatics includingnon-aromatics fed to the dealkylation and conversion products of thedealkylation. Lights such as methane, ethane, propane, and the like thathave been removed from the phenyl ring can be removed from the benzenein the benzene product stream by distillation, including flashdistillation, or by liquid gas separation where the conditions of theproduct stream are not adequate to maintain the lights in the liquidstream.

The detergent range olefin-containing fraction, if the dehydrogenatedproduct has not already been subjected to selective hydrogenation, isselectively hydrogenated and is used for the alkylation of benzene.

Various processes have been proposed for the alkylation of benzene. See,for instance, Peter R. Pujado, Linear Alkylbenzene (LAB) Manufacture, inHandbook of Petroleum Refining Processes, Second Edition (1997), at pp.1.53 to 1.66. The most common processes are those involving the presenceof hydrogen fluoride (“HF Processes”) and those involving the use of asolid acidic catalyst (“Solid Catalyst Processes”). In general, the HFProcess and the Solid Catalyst Process involve contacting an olefin witha stoichiometric excess of benzene at elevated temperature to producealkylbenzene.

The amount of benzene present during the alkylation will depend upon thesought conversion selectivity to alkylbenzene. In general, the greaterthe stoichiometric excess of benzene, the greater the selectivity toalkylbenzene. Typically, the mole ratio of benzene to olefin duringalkylation is within the range of about 5:1 to 50:1 or more. For the HFProcesses, this ratio is often between about 5:1 to 10:1, and for theSolid Catalyst Processes, between about 10:1 to 30:1. Theolefin-containing feed to the alkylation reactor should be sufficientlyfree of impurities, such as water and sulfur compounds that can undulyadversely affect the life of the alkylation catalyst.

U.S. Pat. No. 4,463,205, herein incorporated by reference in itsentirety, discloses typical HF Processes. In these HF Processes, thereactants are normally subjected to vigorous mixing and agitation at thepoint of initial contact of the olefin and the liquid-phase hydrogenfluoride. The alkylation zone preferably has an overall arrangementsimilar to that shown in U.S. Pat. No. 3,494,971. In this arrangement,the two feed hydrocarbons and liquid phase HF are charged to a reactor.The effluent of this reactor is passed into a first settling zone andseparated into HF and hydrocarbon phases. The HF is withdrawn anddivided into a portion passed into a regenerator and a portion returnedto the reactor. A hydrocarbon phase is withdrawn from the first settlingzone and charged to a contactor, which is sometimes referred to as asecond “reactor” as the only hydrocarbon charged to the contactor. TheHF charged to the contactor is a mixture of newly regenerated HF and HFwithdrawn from a second settling zone, which receives the total effluentof the contactor. A portion of the HF withdrawn from the second settlingzone is charged to the reactor to replace the HF withdrawn forregeneration. The hydrocarbon phase which is withdrawn from the secondsettling zone may be withdrawn as the alkylation zone effluent streambut is preferably passed into a stripping column in which dissolved HFis removed overhead and some of the feed aromatic hydrocarbon is alsorecovered. The net bottoms stream of this HF stripping column becomesthe alkylation zone effluent stream charged to the fractionation zone ofthe subject process.

The alkylation reaction zone is maintained at alkylation-promotingconditions. As used herein, the term “alkylation-promoting conditions”is intended to include a pressure sufficient to maintain the reactantsand HF in a liquid phase. A general range of operating pressures is fromabout 200 kpa(a) (29.0 psi(a)) to 4500 kPa(a) (653 psi(a)). Thetemperature range covered by this set of conditions is from about −20°C. (−4° F.) to about 95° C. (203° F.), but the reaction is preferablyconducted at a temperature of from 15° C. (59° F.) to 70° C. (158° F.).The volumetric ratio of HF to the total amount of hydrocarbons enteringthe reactor should be maintained within the broad range of from about0.2:1 to about 10:1, preferably from 0.5:1 to 2:1.

The effluent streams leaving the reactor will typically be an intimateadmixture of liquid phase hydrocarbons and liquid phase hydrogenfluoride. They may be in the form of a true emulsion. A considerableresidence time is normally required to separate these two liquid phases,and the effluent streams are therefore passed into settling zones. Thetwo settling zones will normally be maintained at a temperature that isset by the entering HF-hydrocarbon mixtures withdrawn from therespective upstream zones. They will therefore be at substantially thesame temperature as the immediately upstream reactor or contactor. Thesame is also normally true for the pressures used in the settling zonesafter adjustment for any pressure change due to liquid flow andelevation differences. The settling zones may however be downstream ofcontrol valves and therefore operated at a somewhat reduced pressure.This reduced pressure, however, must be superatmospheric and sufficientto maintain liquid phase conditions. A residence time for both the acidand hydrocarbon phases in the settling zones should be in excess of 30seconds but less than 30 minutes.

Regeneration of the HF catalyst is normally accomplished by strippingthe acid under conditions sufficient to decompose alkylfluorides and toproduce an overhead vapor stream containing HF and the stripping media.Benzene available within the process is a suitable stripping media. Theoverhead vapor stream of the column is passed into a condenser, theresultant condensate is allowed to separate into an acid phase and abenzene phase containing dissolved HF.

The overhead vapor stream of the HF regeneration column will containvapor-phase HF and the benzene. This vapor stream has a lowconcentration of the higher boiling impurities. The higher boilingmaterials are concentrated into a relatively small stream removed fromthe HF regeneration column as a net bottoms stream. The benzene presentin the overhead vapor stream of the HF regeneration column is derivedmainly from the reflux liquid fed to the top of this column. A smallamount of benzene is also dissolved in the liquid phase HF stream fed toHF regeneration column. The reflux liquid is preferably removed from theoverhead receiver, which collects the condensed overhead of the HFstripping column. It is not necessary to supply reflux liquid forsuccessful operation of the HF regeneration column if the feed stream ispassed into the top of the column.

The hydrocarbonaceous phase removed from the second settling zone ispreferably passed into a stripping column referred to as the HFstripping column. The function of this column is to prevent the passageof HF into the downstream fractionation zone. Representative conditionsfor the operation of the HF stripping column include an overhead vaportemperature of about 100° C. (212° F.) to 125° C. (257° F.) at apressure of about 200 kPa(g) (29.0 psi(g)) to 300 kPa(g) (43.5 psi(g)).There is normally no external reflux to this column. The overhead vaporstream of the HF stripping column is normally condensed by cooling it toabout 35° C. (95° F.) or less.

In the Solid Catalyst Processes, benzene and the olefin are reactedunder alkylation conditions in the presence of a solid alkylationcatalyst. These alkylation conditions generally include a temperature inthe range between about 80° C. (176° F.) and about 200° C. (392° F.),most usually at a temperature not exceeding about 175° C. (347° F.).Since the alkylation is typically conducted in at least partial liquidphase, and preferably in either an all-liquid phase or at supercriticalconditions, pressures must be sufficient to maintain reactants in thedesired phase. The requisite pressure necessarily depends upon theolefin and temperature, but normally is in the range of about 1300kPa(g) (189 psi(g)) to 7000 kPa(g) (1015 psi(g)), and most usuallybetween about 2000 kPa(g) (290 psi(g)) and 3500 kPa(g) (508 psi(g)).

Alkylation of benzene by the olefins is conducted in a continuousmanner, and one or more catalyst beds in flow series are used. Forpurposes herein, a catalyst bed is termed a reactor whether in the sameor a separate vessel from another bed. In preferred process embodiments2 or 3 reactors can be used to achieve an advantageous combination ofperformance and capital expense avoidance.

The catalyst may be used as a packed bed or a fluidized bed. The feed tothe reaction zone may be passed either upflow or downflow, or evenhorizontally as in a radial bed reactor. In one desirable variant,olefin-containing feedstock may be fed into several discrete pointswithin the reaction zone, and at each zone the benzene to olefin moleratio may be greater than 50:1. The total feed mixture, that is,aromatic compound plus olefin, is often passed through the packed bed ata liquid hourly space velocity (LHSV) between about 0.3 and about 6 hr⁻¹depending upon, e.g., alkylation temperature and the activity of thecatalyst. Lower values of LHSV within this range are preferred.

Any suitable alkylation catalyst may be used in the process embodiments,provided that the requirements for conversion, selectivity, and activityare met. U.S. Pat. Nos. 5,196,574 and 5,344,997 describe alkylation ofaromatics using a fluorided silica-alumina catalyst. U.S. Pat. No.5,302,732 describes alkylation of aromatics using an ultra-low sodiumsilica-alumina catalyst. The teachings of U.S. Pat. Nos. 5,196,574,5,302,732, and 5,344,997 are incorporated herein by reference. Preferredalkylation catalysts comprise zeolites having a zeolite framework typeselected from the group consisting of FAU, BEA, MOR, MTW, MWW and NES.Preferred zeolites include mordenite, ZSM-4, ZSM-12, ZSM-20, offretite,gmelinite, beta, NU-87, and gottardiite. Further discussion ofalkylation catalysts can be found in U.S. Pat. Nos. 6,315,964 and6,617,481.

In the preferred embodiments, the alkylation conditions regardless ofwhether the HF Process or Solid Catalyst Processes are used, are suchthat the alkylation effluent contains less than about 1000, preferablyless than about 100, ppm-mass of unreacted olefin.

As mentioned above, the alkylation reaction is exothermic. One methodfor assisting in the abatement of any exotherm is to feed a relativelydilute olefin-containing stream to the alkylation. Benzene, by beingsupplied in a greater than stoichiometric amount, serves as a heat sinkas well as kinetically driving the reaction to the mono-alkylbenzenes.Where the detergent range olefin has been derived by dehydrogenating aparaffin of similar carbon number, the olefin-containing feed to thealkylation is about 85 to 90 mass-percent paraffin. In certain of thepreferred embodiments, the detergent olefin fed to the alkylation stephas little, if any paraffin. In such event, the benzene to olefin ratiomay be increased to provide the desired heat sink. Alternatively, astream containing C₅ and C₆ paraffins which is essentially devoid ofolefins may be supplied to the alkylation in an amount sufficient toprovide the sought heat sink. Often, the mole ratio of the C₅ and C₆paraffins in this stream to the detergent range olefins in the feed tothe alkylation is between about 1:1 to 20:1.

The paraffins, having a similar boiling point to benzene, can berecovered from the alkylation effluent with the benzene and recycledwith the benzene. To prevent build-up of these paraffins, a purge may betaken from the benzene loop. The purge may be used for fuel value butpreferably the hydrocarbon values are recovered. For instance, benzenemay be recovered from the purge by extraction, membrane separation orsorption. At least a portion of the recovered benzene may be returned tothe alkylation. At least a portion of the paraffin-containing fractionfrom the separation may be returned to the dehydrogenation zone.Alternatively, at least a portion of the purge without separation ofbenzene may be returned to the dehydrogenation step. The benzene canthen be recovered from a purge from the dehydrogenation-chain grownreaction loop as discussed above.

The alkylation reactor effluent is passed to a distillation assemblythat separates benzene for recycle. If significant amounts of paraffinsin the C₈ and higher range are present, it may also be necessary toprovide a further distillation to separate as the overhead, paraffins,and the bottoms from this second distillation assembly would be fed to aheavies distillation assembly where the alkylbenzene product iscontained in the overhead. If desired, a finishing column may be used tofurther purify the alkylbenzene, especially after a clay treatment toremove color precursors. In this type of distillation train, the bottomsstream for the lights distillation is normally fed to the distillationassembly for separating the benzene. If the alkylation effluent containsunreacted olefin, then the use of the second distillation assembly ispreferred to remove the olefin.

Typically the benzene distillation is generally conducted with a bottomstemperature of less than about 300° C. (572° F.), preferably less thanabout 275° C. (527° F.), usually between about 230° C. (446° F.) and270° C. (518° F.), and at a pressure at which the overhead is providedof between about 5 kPa(g) (0.7 psi(g)) and 300 kPa(g) (43.5 psi(g)),preferably between about 35 kpa(g) (5.1 psi(g)) and 70 kPa(g) (10.2psi(g)). Where a lower carbon number paraffin is present, e.g., as aheat sink, the paraffin may also be contained in the benzene overhead.Generally, the distillation assembly is operated to assure that thebottoms stream contains virtually no benzene. Preferred processesprovide additional advantages with respect to the benzene column. First,where the feed to the alkylation is substantially devoid of C₈ andhigher paraffins, the separation can be effected with fewer distillationplates and less reboiler energy consumption than a conventionalalkylbenzene process where benzene must be separated from paraffinscorresponding in carbon number to the detergent range olefins. Second,since the aromatics that are co-produced with the olefins in thedehydrogenation can be readily separated from the detergent range olefinby distillation, the need for a sorption to remove aromatics from thedetergent range olefins is obviated. As the sorption uses highly purebenzene for regeneration, the need for such highly pure benzene isobviated. Thus additional capital and operating costs in the refining ofthe alkylbenzene can be avoided.

The benzene distillation assembly may comprise one or more distillationcolumns. More than one overhead may be obtained from the benzenedistillation assembly. Each column may contain any convenient packing ordistillation trays, but most often trays such as sieve and bubble trays,are used. Often the assembly provides at least about 3, say 5 to 70, andpreferably 5 to 10, theoretical plates. The reflux ratio (herein definedas the distillate to reflux mass ratio) is often in the range of about2:1 to 1:10, preferably about 1.5:1 to 1:5. The bottoms stream from thebenzene distillation generally contains less than about 1000 parts bymillion by mass (mppm), preferably less than about 50 mppm, andsometimes less than about 5 mppm, benzene. The benzene distillation mayoccur in a single column or two or more distinct columns may be used.For instance, a stripping column may be used to remove a portion, e.g.,20 to 50 percent, of the benzene and then the bottoms from the strippingcolumn would be subjected to rectification in a subsequent column toobtain the desired separation.

If necessary, a paraffin distillation is conducted. The paraffinsdistillation typically involves a bottoms temperature of less than about300° C. (572° F.), preferably less than about 275° C. (527° F.), usuallybetween about 250° C. (482° F.) and 275° C. (527° F.), and a pressure atwhich overhead is provided of between about 5 kPa(a) (0.7 psi(a)) and 10kPa(a) (16 psi(a)), preferably between about 10 kPa(a) (1.5 psi(a)) and50 kPa(a) (7.3 psi(a)). The column may contain any convenient packing ordistillation trays, but most often sieve trays are used. Often theparaffins distillation assembly provides at least about 5, say 7 to 20,theoretical plates. The reflux ratio is often in the range of about 3:1to 1:10, preferably about 1:1 to 1:3. The bottoms stream from theparaffins distillation generally contains less than about 5000,preferably less than about 500, mppm paraffins and preferably less thanabout 10, often less than about 1, mppm benzene. The overhead from theparaffins column generally contains less than about 1, preferably lessthan about 0.5, and most preferably less than about 0.2, mass-percentalkylbenzene. The paraffins distillation may occur in a single column ortwo or more distinct columns may be used.

In the preferred aspects of this embodiment, no paraffins distillationis necessary and the bottoms from the benzene distillation assembly isdirectly provided to the heavy alkylate distillation.

The heavy alkylate distillation is generally conducted with a bottomstemperature of less than about 300° C. (572° F.), preferably less thanabout 275° C. (527° F.), usually between about 250° C. (482° F.) and275° C. (527° F.), and at a pressure of between about 0.5 kPa(a) (0.07psi(a)) and 30 kPa(a) (4.4 psi(a)), preferably between about 1 kPa(a)(0.14 psi(a)) and 5 kPa(a) (0.7 psi(a)). The column may contain anyconvenient packing or distillation trays, but most often structuredpacking is used. Often the heavy alkylate distillation assembly providesat least about 5, say 10 to 30, and preferably 10 to 20, theoreticalplates. The reflux ratio is often in the range of about 2:1 to 1:5,preferably about 0.2:1 to 1:1. The overhead from the heavy alkylatedistillation generally contains less than about 1000 mppm, preferablyless than about 100 mppm, and sometimes less than about 50 mppm, totalheavies.

The refining system may contain additional distillation zones, e.g., torecover additional alkylbenzene from heavies. At least a portion of theheavies-containing stream may be recycled to the dealkylation processfor recovery of benzene.

Where an alkylbenzene stream contains olefinic components, the bottomsfrom the benzene distillation assembly or the overhead from the heavyalkylate distillation column can be subjected to a catalytic operationto reduce bromine index. In this catalytic operation, the alkylbenzenestream is passed to a catalytic conversion zone containing an acidiccatalyst under olefin reduction conditions. The particular unitoperation is not critical to the broad aspects of the processembodiments and any suitable operation may be used.

A number of processes for improving the quality of alkylbenzenes andreducing olefin content have been proposed. The catalysts may be clay ormolecular sieve (natural or synthetic). Included in the clays aremontmorillonite, laponite, saponite, and pillared clays. Filtrol F-24(Engelhard Corporation, Iselin, N.J.) is a preferred clay. Molecularsieves include zeolites A, beta, L, S, T, X and Y and omega, mordenite,erionite, chabazite, boggsite, cloverite, gmelinite, offretite, pentasilmolecular sieves having silica to alumina mole ratios greater than about10, and silicoaluminophosphates (such as SAPO-5 and SAPO-41).

The olefin reduction is typically conducted at temperatures betweenabout 25° C. (77° F.) and 250° C. (482° F.), and most often betweenabout 70° C. (158° F.) and 150° C. (302° F.), under a pressuresufficient to maintain the alkylbenzene stream under liquid conditions,e.g., within the range of about 0.1 kPa(a) (0.01 psi(a)) to 150 kPa(a)(21.8 psi(a)). The contact time with the catalyst is sufficient toprovide the desired reduction in bromine index. For a fixed bed system,the weight hourly space velocity is typically in the range of about 0.1to 20 hr⁻¹. The bromine index of the treated alkylbenzene stream ispreferably below about 10, more preferably below about 2. The olefinreduction conditions also cause byproducts such as the formation ofdialkylbenzene and benzene from alkylbenzene and form oligomers andpolymers from olefinic components.

The effluent from the olefin reduction is subjected to the fourthdistillation to remove as an overhead, benzene, and heavies such as thedialkylbenzene and the oligomers and polymers from olefinic components.

With reference to FIG. 1, an apparatus is depicted that is capable ofusing dual feedstocks for making alkylbenzenes. In apparatus 100, aC₅-C₆-containing feedstock is provided via line 102 to oxygenateextractor 104. This feedstock may be obtained from any suitable sourceincluding a naphtha-range fraction or a Fischer-Tropsch process and thususually contains normal paraffins. For purposes of this discussion, thefeedstock contains olefins and oxygenates as well as some C₅-C₆paraffins and some higher and some lower molecular weight paraffins andolefins. A methanol-water sorbent is introduced into extractor 104 vialine 106 and the spent sorbent is withdrawn via line 108 forregeneration.

The C₅-C₆-containing feedstock having oxygenates removed is directedfrom extractor 104 to dimerization reactor 112 via line 110. While adimerization reactor is shown, reactor 112 could also represent ametathesis reactor. A portion of the olefin in the C₅-C₆-containingfeedstock is consumed to provide dimerized olefin primarily in the C₁₀to C₁₂ range with some higher and some lower molecular weight olefins.The olefin will have some branching, but typically contains very littlequaternary carbon branching. The presence of paraffins in the feedstockto the reactor advantageously serves as a heat sink to facilitatetemperature control during the exothermic dimerization reaction. Thefeedstock now containing the dimer is passed from reactor 112 via line114 to fractionation assembly 116 where lights such as hydrogen andcracking products are removed via line 120 and a stream containing C₅and C₆ aliphatic hydrocarbons (paraffins and unreacted olefins) iswithdrawn via line 122 and passed to dehydrogenation reactor 124 toconvert a portion of the paraffins to the corresponding olefins. As thedehydrogenation is equilibrium affected, the product will containolefins and paraffins. The dehydrogenation product is passed via line126 to dimerization reactor 112 where in combination with the feedstock,the olefins are subjected to dimerization reaction conditions.

As discussed above, by directing the olefin-containing feedstock firstto the dimerization reactor, enhanced efficiencies in thedehydrogenation and dimerization cycle loop are achieved since the feedto the dehydrogenation contains a lower concentration of olefin thandoes the C₅-C₆-containing feedstock.

The dimerized olefin is withdrawn from fractionation assembly 116 vialine 128 and is introduced into selective hydrogenation reactor 130where diolefins are converted to mono-olefins. Hydrogen required for theselective hydrogenation is supplied by line 131. The effluent fromreactor 130 which contains olefins having a reduced concentration ofdienes, is passed via line 132 to stripper 134 for removal of light endscontained in an overhead by line 136. Depending on the hydrogen and theC₅-C₆ contents of the overhead, the overhead may be recovered andrecycled to the selective hydrogenation reactor or the C₅-C₆dehydrogenation reactor.

Selective hydrogenation reactor 130 also treats a dehydrogenationproduct generated from a C₁₀-C₁₂ paraffin-containing feedstock suppliedby line 138. This feedstock is dehydrogenated in dehydrogenation reactor140 and the dehydrogenation effluent from reactor 140 is directed byline 142 to separator 144. The overhead from separator 144 is exhaustedvia line 146 and contains hydrogen. A portion of the overhead isrecycled via line 148 to dehydrogenation reactor 140. The bottoms fromseparator 144 are passed via line 150 to selective hydrogenation reactor130.

The bottoms from stripper 134 are passed via line 152 to sorption system154 in which aromatics are selectively removed from the selectivehydrogenation effluent. A suitable sorption system is described in, forinstance, U.S. Pat. No. 5,276,231 and U.S. Pat. No. 5,334,793, hereinincorporated by reference in its entirety. Aromatics are formed in thedehydrogenation of the C₁₀-C₁₂-paraffin containing feedstock. As asorption system is provided, the conditions of dehydrogenation reactor124 may include those in which benzene and other aromatics are formed.These more severe conditions tend to increase the conversion ofparaffins to olefins per pass through dehydrogenation reactor 124.

The olefin-containing effluent from sorption system 154 is passed vialine 156 to alkylation reactor 160 for reaction with benzene supplied byline 166 to produce alkylbenzenes. The bottoms from benzenefractionation assembly 164 contain alkylbenzenes and paraffins. Usuallythe alkylation reactor 160 is operated under conditions such thatessentially all of the olefins are reacted. As at least a portion of thedetergent range olefin for alkylation is sourced from a C₅-C₆-containingfeedstock, the olefin-containing feed will contain less, and perhaps insome instances, essentially no, similar molecular weight paraffin.Accordingly, the reboiler heat required for recovery of the alkylbenzeneproduct from the alkylation reactor effluent can be appreciably reduced.

The alkylation reaction product is directed by line 162 to benzenefractionation assembly 164 with benzene being recovered as an overheadand recycled to alkylation reactor 160 through line 166. Benzene make-upis provided to benzene fractionation system 164 by line 170. The bottomsstream from benzene fractionation assembly 164 is fed to paraffinsfractionation assembly 174 via line 172. The overhead from paraffinsfractionation assembly 174 contains paraffins that are recycled todehydrogenation reactor 140 via line 176. The bottoms fraction fromparaffins fractionation assembly 174 is directed via line 178 to heaviesfractionation assembly 180 with the purified alkylbenzene beingwithdrawn as the overhead via line 182 and a heavies-containing bottomsstream being withdrawn via line 184. The heavies-containing bottomsstream may, if desired, be subjected to an additional fractionation torecover additional alkylbenzenes.

Turning to FIG. 2, the depicted apparatus uses only a C₅-C₆-containingfeedstock as described in connection with FIG. 1. In apparatus 200, aC₅-C₆-containing feedstock is provided via line 202 to oxygenateextractor 204. This feedstock may be obtained from any suitable sourceincluding a naphtha-range fraction or a Fischer-Tropsch process. Forpurposes of this discussion, the feedstock is a naphtha range feedstockand contains some higher and lower molecular weight paraffins. Amethanol-water sorbent is introduced into extractor 204 via line 206 andthe spent sorbent is withdrawn via line 208 for regeneration.

The C₅-C₆-containing feedstock having oxygenates removed is directedfrom extractor 204 to dehydrogenation reactor 212 by line 210. Thedehydrogenation reactor effluent is then passed via line 214 tometathesis reactor 216 to provide a metathesis reaction productincluding olefins having from 10 to 12 carbon atoms per molecule.Usually the reaction product contains very little quaternary carbonbranching. Advantageously, the metathesis reaction produces little, ifany, aromatics. The metathesis product is passed to distillationassembly 220 via line 218. The sought C₁₀ to C₁₂ olefins are containedin the bottoms stream. An overhead containing hydrogen and low molecularweight hydrocarbons is exhausted via line 222 and the remainingcomponents, primarily paraffins and lower molecular weight olefins, arerecycled from distillation assembly 220 to dehydrogenation reactor 212via line 224.

The bottoms stream is passed through line 226 to selective hydrogenationreactor 228 where diolefins are converted to mono-olefins. The effluentfrom reactor 228 which contains olefins having a reduced concentrationof dienes, is passed via line 230 to stripper 232 for removal of alight-ends-containing overhead by line 234. Depending on the hydrogenand the C₅-C₆ contents of the overhead, the overhead may be recoveredand recycled to selective hydrogenation reactor or the dehydrogenationreactor. Any additional hydrogen required for the selectivehydrogenation is supplied by line 236.

Because the olefin-containing feed is essentially devoid of aromatics,the sorption system as described in connection with FIG. 1 is notnecessary. Hence, the bottoms fraction from stripper 232 that treats theselective hydrogenation product, is directly fed via line 238 toalkylation reactor 240 for reaction with benzene supplied by line 248.

The alkylation reaction product is directed by line 242 to benzenefractionation assembly 244 with benzene being recovered as an overheadand recycled to alkylation reactor 240 through line 248. Benzene make-upis provided to benzene fractionation system 244 by line 250. The bottomsstream from benzene fractionation assembly 244 is fed via line 252 toheavies fractionation assembly 254 with the purified alkylbenzene beingwithdrawn as the overhead via line 256 and a heavies-containing bottomsstream being withdrawn via line 258. The heavies-containing bottomsstream may, if desired, be subjected to an additional fractionation torecover additional alkylbenzenes. As the olefins can beessentially-completely consumed in the alkylation reaction and theolefins fed to alkylation reactor 240 contain little, if any, paraffin,the necessity of a paraffins fractionation assembly is obviated. To theextent that paraffins are contained in the olefin feed stream toalkylation reactor 240, they can be passed to the overhead of benzenefractionation assembly 244. A purge can prevent build-up of theparaffins in the alkylation reactor-benzene fractionation assembly loop.

In FIG. 3, the depicted apparatus uses an olefin and paraffin-containingfeedstock having predominantly hydrocarbons in the C₅-C₆ range. Inapparatus 300, the C₅-C₆-containing feedstock is provided via line 302to oxygenate extractor 304. This feedstock may be obtained from anysuitable source including a naphtha-range fraction or a Fischer-Tropschprocess. For purposes of this discussion, the feedstock is a naphtharange feedstock and contains some higher and lower molecular weightparaffins and olefins. A methanol-water sorbent is introduced intoextractor 304 via line 306 and the spent sorbent is withdrawn via line308 for regeneration.

The C₅-C₆-containing feedstock having oxygenates removed is directed byline 310 from extractor 304 to selective hydrogenation unit 328 wherediolefins are converted to mono-olefins. As will be discussed later,dehydrogenated product is also passed to selective hydrogenation unit328. Selective hydrogenation unit 328 has a reactor and a stripper forremoval of light ends in a stripper overhead. The stripper overhead maybe recovered and recycled to the selective hydrogenation reactor or thedehydrogenation reactor, depending on the hydrogen and the C₅-C₆contents of the stripper overhead. Make-up hydrogen required for theselective hydrogenation is supplied by line 336. Selective hydrogenationunit 328 removes dienes and any alkynes from the feedstock and from therecycling dehydrogenated product. Advantageously, the olefins in thefeedstock are used for the dimerization. By not introducing thefeedstock immediately before the dehydrogenation, the olefins in thefeedstock do not adversely affect the amount of conversion of paraffinto olefin in the dehydrogenation. This efficient use of valuable olefinsin the feedstock is economically advantageous.

The effluent from reactor 328 which contains olefins having a reducedconcentration of diolefins is passed via line 330 to dimerizationreactor 312 to provide a dimerization reaction product includingdetergent range olefins. The dimerization reactor effluent is passed todistillation assembly 320 via line 314. The sought detergent rangeolefins are contained in a higher boiling fraction withdrawn via line338. A bottoms stream containing trimers and highers is removed via line326. An overhead containing hydrogen and low molecular weighthydrocarbons is exhausted via line 322 and the remaining components,primarily paraffins, aromatics and lower molecular weight olefins, arewithdrawn as a side stream and passed via line 318 to dehydrogenationreactor 316. A dehydrogenated product is withdrawn from dehydrogenationreactor 316 via line 324 and recycled to selective hydrogenation unit328.

The aromatics are formed during the dehydrogenation. Benzene is theprimary aromatic formed, however, some toluene and other alkylaromaticsmay be formed. A purge is taken via line 360 to maintain the aromaticsrecycling to the dimerization reactor at a desired concentration. Thispurge in line 360 is passed to catalytic dealkylation reactor 362 wheretoluene and other alkylaromatics are dealkylated to form benzene. Thedealkylation is conducted under a hydrogen atmosphere. Duringdealkylation, olefins are hydrogenated. The dealkylated product ispassed via line 364 to stripper 368 where an overhead containing lightsis withdrawn via line 370 and a bottoms stream containing benzene andparaffins is passed via line 372 to alkylation reactor 340 as discussedfurther below.

The detergent range olefin fraction from distillation assembly 320 ispassed through line 338 to alkylation reactor 340. Because the detergentrange olefin fraction is essentially devoid of aromatics, the sorptionsystem as described in connection with FIG. 1 is not necessary. Hence,the fraction is directly fed to alkylation reactor 340 for reaction withbenzene supplied by lines 348 and 372. Advantageously, the paraffinsco-supplied with benzene via line 372 serve as a heat sink for theexothermic alkylation reaction.

The alkylation reaction product is directed by line 342 to benzenefractionation assembly 344 with benzene being recovered as an overheadand recycled to alkylation reactor 340 through line 348. Benzene make-upis provided to benzene fractionation system 344 by line 350. Since aportion of the benzene for the alkylation reaction is provided bydehydrogenating the feedstock, the amount of benzene make-up providedvia line 350 can be significantly reduced. To maintain the paraffins ata desired concentration in alkylation reactor 340, a purge is taken fromline 348 by line 374. This purge will be rich in benzene and containparaffins. As shown, line 374 directs the purge to membrane separator376 which has a membrane selective for the permeation of acyclicparaffins as compared to benzene. The paraffin-rich permeate is passedvia line 378 to distillation assembly 320 such that paraffins can bepassed to dehydrogenation reactor 316. A retentate fraction containingbenzene is passed via line 380 to alkylation reactor 340. Alternatively,the entire purge stream may be recycled to dehydrogenation reactor 316,or it may be subjected to selective sorption to remove aromatics.

The bottoms stream from benzene fractionation assembly 344 is fed vialine 352 to heavies fractionation assembly 354 with the purifiedalkylbenzene being withdrawn as the overhead via line 356 and aheavies-containing bottoms stream being withdrawn via line 358. Theheavies-containing bottoms stream may, if desired, be subjected to anadditional fractionation to recover additional alkylbenzenes. As theolefins can be essentially-completely consumed in the alkylationreaction and the olefins fed to alkylation reactor 340 contain little,if any, paraffin, the necessity of a paraffins fractionation assembly isobviated. To the extent that paraffins are contained in the olefin feedstream to alkylation reactor 340, they can be passed to the overhead ofbenzene fractionation assembly 344. A purge can prevent build-up of theparaffins in the alkylation reactor-benzene fractionation assembly loop.

1. An integrated process for producing detergent range alkylbenzenesfrom a feedstock predominantly comprising C₅ and C₆ paraffins andcontaining C₅ and C₆ mono-olefins, the process comprising: a) subjectingthe feedstock to chain growth reaction conditions for converting C₅ andC₆ mono-olefins to C₁₀ to C₁₂ mono-olefins to provide a reaction productcontaining C₅ and C₆ paraffins and C₁₀ to C₁₂ mono-olefins; b)subjecting at least a portion of the reaction product to distillation toprovide a higher boiling fraction containing C₁₀ to C₁₂ mono-olefins anda lower boiling fraction comprising C₅ and C₆ paraffins; c) subjectingat least a portion of the lower boiling fraction of step b todehydrogenation conditions to provide a dehydrogenated productcomprising C₅ and C₆ mono-olefins; d) recycling at least a portion ofthe dehydrogenated product to step a; e) alkylating benzene with atleast a portion of the higher boiling fraction of step b underalkylation conditions including a stoichiometric excess of benzene toolefin to provide an alkylation effluent comprising alkylbenzenes andbenzene; and f) separating benzene from at least a portion of thealkylation effluent and recycling at least a portion of the separatedbenzene to step e.
 2. The process of claim 1 wherein the feedstockcontains about 2 to 10 mass-percent C₅ and C₆ mono-olefins.
 3. Theprocess of claim 2 wherein the feedstock is derived from a naphtha rangefraction.
 4. The process of claim 3 wherein the feedstock is derivedfrom a Fischer-Tropsch synthesis.
 5. The process of claim 2 wherein thechain growth reaction of step a is a dimerization.
 6. The process ofclaim 2 wherein the chain growth reaction of step a is a metathesis. 7.The process of claim 2 wherein the at least a portion of the higherboiling fraction of step b contains less than about 0.5 mass-percent ofparaffins having 9 or more carbon atoms per molecule.